Catalytic cracking of hydrocarbon feedstocks has been characterized by certain basic steps repeated in cyclic manner. The catalysts are generally highly porous solids characteristically having extensive surface area and possessing acidic sites. During a relatively short period of time, hydrocarbon charge, such as gas-oil, undergoes profound conversions of a complex nature on contact with those surfaces at elevated temperatures upwards of about 850.degree. F. and essentially atmospheric pressure. The temperatures may range up to about 1000.degree. F. and the pressure on incoming charge is usually only enough to overcome pressure drop through the reactor and associated produce recovery facilities, say 30 to 50 psig.
The conversions taking place in the presence of the cracking catalyst include scission of carbon-to-carbon bonds (simple cracking), isomerization, polymerization, dehydrogenation, hydrogen transfer and others leading to lighter, lower molecular weight compounds as important desired products. In many installations motor gasoline of end point near about 400.degree. F. is a primary product and cracking units are often operated to maximize high quality gasoline within constraints imposed by ability to profitably market the unavoidable by-products such as butane and lighter. In addition to the gaseous by-products, the reactions on cracking catalyst also produce hydrocarbons of very low volatility and very high carbon content which remain on the active surfaces of the catalyst and mask the active sites, rendering the catalyst inactive. Those deposits of heavy carbonaceous matter (commonly called "coke") can be removed by burning with air to restore the active surface and thus regenerate the activity of the catalyst. Typically the coke contains about 6-7% hydrogen on a weight basis. In commercial plants for practice of catalytic cracking, catalyst inactivated by coke is purged of volatile hydrocarbons, as by steaming, and contacted with air at elevated temperature to burn off the coke.
Combustion of the coke generates carbon dioxide, carbon monoxide and water as combustion products and releases large amounts of heat. To a very considerable extent, the heat so released has been applied to supply the endothermic heat of reaction during the cracking phase of the cycle. In its earliest stages, catalytic cracking was conducted in fixed beds of catalyst provided with exchange tubes through which a heat transfer fluid was circulated to abstract heat during regeneration and supply heat during cracking. Continuity of operation was achieved by a complex system of manifolds and valves serving a plurality of reactors such that one is used for cracking while two or more others were purged of volatiles, regenerated, again purged and ready to assume the cracking function as catalyst in the first reactor became spent.
Further development made available systems in which the catalyst is moved continuously through a reactor, purged, transferred to a regenerator, again purged and returned to the reactor. These moving catalyst systems are able to dispense with the circulating heat transfer medium and instead employ the catalyst itself as a medium for conveying heat from the regenerator to the reactor. The early catalysts such as acid-treated clays, and synthetic amorphous silica-alumina composites, resulted in deposition of quantities of coke in excess of the amounts which on complete combustion to carbon dioxide and water will supply the heat of reaction required by the reactor. In some installations, a portion of the heat was withdrawn by heat exchange coils in the regenerator. That practice was followed in the moving compact bed process known as Thermofor Catalytic Cracking (TCC). Another expedient is to circulate a portion of catalyst from the regenerator through a cooling heat exchanger and back to the regenerator. That practice was found suitable for systems in which a finely divided catalyst is suspended in the hydrocarbon charge in the reactor and in the combustion air in the regenerator. These suspended catalyst systems applied the fluidized solids phenomenon and are classed generally as Fluid Catalytic Cracking (FCC).
Characteristic of all the systems for many years was a high content of carbon monoxide (CO) in the flue gas from the regenerator, a result of incomplete combustion or partial utilization of the fuel value of the coke. CO in the flue gas is undesirable for other reasons. That combustible gas can burn in regenerator gas discharge equipment and in flues leading to temperatures which damage those facilities. The loss of potential fuel value has been avoided by providing "CO boilers" in which the CO is burned in contact with steam generation tubes, thus recovering sensible heat from the flue gas as well as fuel value of the CO.
As designs of moving catalyst systems for charging heavier stocks were developed, the cracker received some hydrocarbons in liquid form, requiring heat input for vaporization of charge, heating the charge to reaction temperature and for endothermic heat of reaction. The "heat balanced" FCC design aids in satisfaction of these requirements. Typically, that design provides a heat sensor in the reacted vapors before removal from the reactor. An automatic control of the valve in the line for return of hot regenerated catalyst from the regenerator to the reactor assures return of that amount of hot regenerated catalyst which will maintain reactor top temperature at a desired set point. It will be seen that this control also sets an important reactor parameter, namely ratio of catalyst to oil (C/O), corresponding to space velocity in fixed bed reactors. It follows that, for a given set of regenerator conditions, C/O is a dependent variable not subject to independent control by the operator.
The advent of zeolite cracking catalyst as described in U.S. Pat. No. 3,140,249 introduced new considerations in catalytic cracking design and practice. Such catalysts are highly active, inducing more profound conversion of hydrocarbon charge stock than the older catalysts. In addition, they are some selective in that a larger proportion of the conversion products are motor gasoline components with lesser proportions of gas and coke. Because of that increased selectivity, the zeolite cracking catalyst rapidly became the catalyst of choice, particularly in areas of high gasoline demand, such as the United States. The more active catalyst has been effectively applied in FCC Units at short catalyst contact times, such as the modern riser reactor units in which hot catalyst is dispersed in a column of charge rising through a conduit to an enlarged catalyst disengaging zone. Contact times of 20 seconds or less are common practice in such units. Such short contact times place a premium on high activity of the catalyst. Since activity of the regenerated catalyst is a function of residual coke remaining on the catalyst after regeneration, it becomes important to reduce residual coke to the lowest level economically attainable.
The extent of coke burning is a function of time and temperature. Rate of coke burning increases with increased temperature. In any given installation, the volume of the regenerator imposes a constraint on time of contact between catalyst and regeneration air. Temperature of regeneration is constrained by thermal stability of the catalyst, which suffers unduly rapid loss of activity on exposure to moisture of the regeneration air at temperatures upwards of about 1400.degree. F. In addition, the regeneration temperature must be held to a level which will not cause damage to vessel internals. As regeneration gas rises from a dense bed in a regenerator, burning of CO can take place in a "dilute phase" containing only a small amount of catalyst. Because there is very little catalyst to absorb the heat thus released, the temperature of the gas rises rapidly and may reach levels which cause damage to the cyclones which separate entrained catalyst from regenerator fume, plenum chambers and flues for discharge of the flue gas. This may be combatted by injecting water or steam to these internals.
Better techniques have been recently proposed and adopted in many plants. According to the system of U.S. Pat. No. 3,909,392, catalyst from the dense bed of the regenerator is educted through tubes to the disperse phase, thus providing catalyst mass to absorb heat of CO combustion and return that heat to the dense bed as the catalyst falls back into that bed. A widely practiced technique causes CO combustion to take place in the dense bed by use of a catalyst promoted with platinum or the like in very small amounts. See U.S. Pat. No. 4,072,600. By transferring the heat of burning CO to the dense bed, these developments make higher regeneration temperatures available to regenerate catalyst to lower residual coke levels, hence higher activity.
Regeneration temperatures above 1250.degree. F., preferably around 1300.degree. F. and up to about 1375.degree. F., become feasible at residual coke levels of less than 0.1% by weight on catalyst. The necessary result of regeneration at these increased temperatures is that the automatic control to maintain preset reactor top temperature will reduce the rate of catalyst flow from regenerator to reactor below the rates for lower regeneration temperature, thus reducing C/O. In addition, catalyst at these high temperatures will heat a portion of the charge to excessive levels at which thermal cracking occurs with resultant production of gas, olefins and coke.
Operators of FCC Units have also been concerned about emissions of sulfur dioxide and sulfur trioxide (SO.sub.x) in the regenerator flue gas. The hydrocarbon feeds processed in commercial FCC units normally contain sulfur. It has been found that about 2-10% or more of the sulfur in a hydrocarbon stream processed in an FCC system is transferred from the hydrocarbon stream to the cracking catalyst, becoming part of the coke formed on the catalyst particles within the FCC cracking or conversion zone. This sulfur is eventually removed from the conversion zone on the coked catalyst which is sent to the FCC regenerator. Accordingly, about 2-10% or more of the feed sulfur is continuously passed from the conversion zone into the catalyst regeneration zone with the coked catalyst in an FCC unit.
In an FCC catalyst regenerator, sulfur contained in the coke is burned, along with the coke carbon, forming primarily gaseous sulfur dioxide and sulfur trioxide. These gaseous sulfur compounds become part of the flue gas produced by coke combustion and are conventionally removed from the regenerator in the flue gas.
It has been shown that SO.sub.x in the regenerator flue gas can be substantially cut back by including in the circulating catalyst inventory an agent capable of reacting with an oxide of sulfur in an oxidizing atmosphere or an environment which is not of substantial reducing nature to form solid compounds capable of reduction in the reducing atmosphere of the FCC reactor to yield H.sub.2 S. Upon such reduction, the sulfur leaves the reactor as gaseous H.sub.2 S and organic compounds of sulfur resulting from the cracking reaction. Since these sulfur compounds are detrimental to the quality of motor gasoline and fuel gas by-products, the catalytic cracker is followed by downstream treating facilities for removal of sulfur compounds. Thus the gaseous fractions of cracked product may be scrubbed with an amine solution to absorb H.sub.2 S which is then passed to facilities for conversion to elemental sulfur, e.g. a Claus plant. The additional H.sub.2 S added to the cracker product stream by chemical reduction in the reactor of the solid sulfur compounds formed in the regenerator imposes little additional burden on the sulfur recovery facilities.
The technology heretofore proposed has involved circulating the SO.sub.x binding agent, the latter either being an integral part of the cracking catalyst particles or taking the form of separate particles having essentially the same fluidization properties as the cracking catalyst. Suitable agents for the purposes have been described in a number of previously published documents. Discussion of a variety of oxides which exhibit the property of combining with SO.sub.x and thermodynamic analysis of their behavior in this regard are set out by Lowell et al., SELECTION OF METAL OXIDES FOR REMOVING SO.sub.X FROM FLUE GAS, IND. ENG. CHEM. PROCESS DES. DEVELOP., Vol. 10, No. 3 at pages 384-390 (1971).
An early attempt to reduce SO.sub.x emission from catalytic cracking units, as described in U.S. Pat. No. 3,699,037, involves adding particles of a Group II metal compound, especially calcium or magnesium oxide, to a cracking unit cycle at a rate at least as great as the stoichiometric rate of sulfur deposition on the cracking catalyst, the additive preferably being injected into the regeneration zone in the form of particles greater than 20 microns. Particle size was chosen to assure a relatively long residence time in the unit. In putting the invention into practice, the Group II metal compound is recycled at least in part between the reactor and the regenerator, the remainder leaving the cycle along with catalyst fines entrained in regenerator flue gas. Subsequently it was proposed to incorporate the alkaline earth metal compound in the cracking catalyst particles by impregnation in order to minimize loss of the sulfur acceptor in the regenerator flue gases. See U.S. Pat. No. 3,835,031. This patent apparently recognizes the need for free oxygen for binding SO.sub.x with a Group II metal oxide since the equations for the reaction taking place in the regenerator is summerized as follows: EQU MgO+SO.sub.2 +1/2O.sub.2 =MgSO.sub.4
Similar use of reactive alumina either as a discrete fluidizable entities or as a component of catalyst particles is described in U.S. Pat. Nos. 4,071, 436; 4,115,250 and 4,115,251. Use of oxidants including platinum or chromium as adjuncts to alumina is suggested in these patents. A related use of cerium oxide on an alumina support is described in U.S. Pat. No. 4,001,375.
In the prior art techniques aforementioned, emphasis was on reversibly reacting sulfur oxides in the flue gas, and doing so while the gases were still in the regenerator. Since the sulfur loaded particles were carried to the reactor to be converted to gaseous hydrogen sulfide under the reducing atmosphere created by the cracking operation, the agents used to bind and then release sulfur were necessarily limited to those inherently capable of doing so under the constraints of temperature and time imposed by the operation of the reactor and the regenerator. Such procedures offer promise as means to reduce SO.sub.x emissions from refineries but they leave much to be desired.
With units operating with high sulfur feedstock, relatively large amounts of sulfur acceptor are needed to accomplish reductions in sulfur oxide levels. This will result in appreciable dilution of the active catalyst in the cracking reaction cycle whether the sulfur acceptor is a part of the catalyst particles or is present as discrete entities circulated with catalyst inventory. A basic limitation is that conditions of time and temperature for operating cyclic cracking units, especially heat balanced FCC units, are geared to maximizing production of desired products and conditions that will favor this result, are by no means those that are optimum for reversibly reacting sulfur oxides in the regenerator and carrying the sulfur back to the reactor for conversion at least in part to hydrogen sulfide. The generally recognized need or desirability of having excess oxygen in the regenerator beyond that needed to burn coke, in order to convert oxides to sulfur of trioxide which will then react with metal oxide to form a sulfate, imposes severe limitations on those refineries that operate regenerators in conventional mode, i.e., without full combustion. Thus, introduction of air (oxygen) to promote SO.sub.x association with metal oxide is basically inconsistent with the operation of a regenerator in conventional incomplete combustion mode.